Selective multistage hydrogenation of hydrocarbons



March 8, 1966 c. w. STREED ET AL 3,239,454

SELECTIVE MULTISTAGE HYDROGENATION OF HYDROCARBONS Filed Jan. 14, 1963Euw 93 m w 5 822 M MY W ATTORNEY United States Patent 3,239,454SELECTIVE MULTISTAGE HYDROGENATION 0F HYDROCARBONS Carl W. Streed,Haddonfield, and Raymond R. Halik,

Pitman, N.J., assignors to Socony Mobil Oil Company,

Inc., a corporation of New York Filed Jan. 14, 1963, Ser. No. 252,268 16Claims. (Cl. 208-210) The present invention relates to a process,preferably of the continuous type, for the selective hydrogenation in 3or more catalytic stages of a normally liquid hydrocarbon mixturecontaining aromatic hydrocarbons, olefins, diolefins, sulfur compoundsand possibly acetylenes.

Many processes for the selective hydrogenation of petroleum hydrocarbonsin a nondestructive manner, that is hydrogen addition with little or nocracking or hydrocracking of the feed, have been advanced over theyears. Interest in this field has intensified in the last decade withthe advent of catalytic hydroforming which has made large supplies ofhydrogen-rich gas available at refineries at a relatively low cost whichmight be charged at least in part against the improvement of naphthas inthe reforming operation.

Selective hydrogenation serves many purposes and the instant inventionis particularly concerned with a preferred embodiment in which anunstable hydrocarbon mixture of the type mentioned above containing ahigh proportion of aromatic hydrocarbons is hydrogenated in at least twostages of increasing severity to prepare a stable product from whichvaluable aromatic hydrocarbons can readily be separated, for example, bysolvent extraction with a solvent such as diethylene glycol. In suchextractions it is realtively easy to separate benzene and other aromaticcompounds from parafiinic hydrocarbons but this is not true ofseparating benzene from aliphatic and cycloaliphatic unsaturatedcomponents, and especially from organic sulfur compounds, in themixture.

To prepare a suitable feed for the solvent extraction of aromatics, itis necessary to convert the organic sulfur compounds to a readilyseparable material such as hydrogen sulfide gas, to saturate theunstable gum forming diolefins and also to saturate the mono-olefinswithout at the same time converting aromatic hydrocarbons intonaphthenes by excessive hydrogenation. Although it is simple to specifythe reactions with hydrogen for obtaining these results, achieving themin commercial practice has been an entirely different matter. There isan increasing demand for the production of aromatic hydrocarbons frompetroleum so that the supplies of these compounds are not restricted tothe current production level of the steel and coking industries. Despitethis demand, prior to the present invention there was still no fullysatisfactory commercial method for the the hydrogenation of suchmixtures of aromatic and unsaturated aliphatic hydrocarbons.

It is not feasible to completely saturate and desulfurize such feedstocks in a single operation because the relatively high temperaturessuitable for hydrodesulfurization also promote the formation of coke andolefin polymers (gums) and may hydrogenate aromatics to naphthenes undercertain conditions. Prior to the present invention even conducting thehydrogenation reactions in several stages to avoid or minimize theaforesaid deficiencies has not been entirely satisfactory by reason ofthe accumulation of polymeric deposits that reduce the activity ofhydrogenation catalysts, thereby requiring frequent regeneration, and inaddition these deposits also plug up piping and other equipment. Notonly thermal polymization but also catalytic polymerization must beminimized as many good hydrogenation and desulfurization catalysts alsocatalize the polymerization of diolefins.

While various techniques are known for at least partially reducingpolymer formation of hydrocarbons at elevated temperatures, neverthelesspolymer formation remains a critical problem in commercial plants forthe selective hydrogenation of charging stocks of the type described.

An object of the present invention is to provide an improved process forthe selective hydrogenation of mixed organic compounds.

Still another object of the invention is to provide an improved processfor the selective, nondestructive hydrogenation of a mixture of aromaticand olefinic hydrocarbons boiling below about 500 F., and preferablybelow about 275, in which contact catalysts are kept on stream forsubstantially longer periods.

A further object of the invention is to provide an improved process forthe selective catalytic hydrogenation of a mixture of aromatic andolefinic hydrocarbons in two or more reactors in which the on-streamperiods of all reactors are lengthened.

Other objects and advantages of the invention will be apparent to thoseskilled in the art upon consideration of the following detaileddisclosure in which all temperatures are expressed in terms of degreesFahrenheit, all proportions in terms of weight and all temperatures ofboiling points or ranges are measured at atmospheric pressure by theA.S.T.M. procedure unless otherwise expressly stated hereinafter.

The present invention is an improved method for the selective,nondestructive hydrogenation of a hydrocarbon liquid boiling below about500, which contains aromatic hydrocarbons, preferably in relativelylarge amounts, and also olefins, diolefins and sulfur compounds. Itcomprises reacting said material substantially in the liquid phase withhydrogen in contact with a porous solid hydrogenation catalyst of highhydrogenation activity and low polymerization activity in an initialhydrogenation zone while controlling the hydrogenating conditionstherein to provide an initial hydrogenation effiuent in which asubstantial amount that may be at least about 35% of the diolefins, andpreferably at least about 50%, have been at least partially saturatedand in which a substantial part of said liquid feed and products thereofremain in the liquid phase. This liquid phase may be at least about 20%and preferably at least about 60% of the liquid feed. In a preferredmodification, the Bromine Number of the normally liquid fraction of saideffluent is also reduced at least about 25% below that of the liquidfeed. After vaporizing liquid in said hydrogenation eflluent, theresulting vapors are passed together with hydrogen through anintermediate hydrogenation zone. It is preferable to effect thevaporization of liquid in said initial effluent and separation of thegaseous phase thereof in the presence of a substantial amount of aliquid flux, for example, an amount of flux equal to at least about 5%,and desirably more than about 10%, of said liquid feed. In theintermediate zone, the gaseous mixture is in contact with a porous,solid hydrogenation catalyst having a high hydrogenation activity and alow polymerization activity at a temperature high enough for olefinsaturation and also substantially higher than the average temperature insaid initial zone under conditions controlled to further hydrogenatesaid vapors, for example to the extent whereby the diolefin content ofthe normally liquid fraction thereof is less than about 50%, andpreferably less than about 40%, of that of said liquid feed. Next, theefiluent is withdrawn from the intermediate Zone and passed through asubsequent conversion zone at a temperature suitable for desulfurizationin contact with a porous, solid, sulfur-resistant conversion catalysthaving at least moderate hydrogenation activity and a highdesulfurization activity while controlling conversion conditions in saidconversion zone to produce an eflluent and preferably less than about 2,and an organic sulfur,

con-tent below about 20 p.p.rn., and preferably below about 15 p.p.m. Ina preferred modification, the temperature at the inlet oftheintermediate hydrogenation zone is also high enough fordesulfurization.

The aforesaid flux may be any liquid substance or mixture which ismiscible with said liquid feed and does not react with it, includingimmediate reaction products, recycled partially or fully hydrogenated.products of the instant process or extraneous liquids, preferably thosehaving a substantial content of aromatic hydrocarbons. The flux may bewithdrawn from said vaporization step and from the process in an amountequal to at least about 0.5% (about 110% being preferred) of said liquidfeed thereby preventing the accumulation of any substantial quantity. ofpolymeric material in the equipment.

7 Narrower aspects of the invention relate to inhibiting thehydrogenation of aromatic hydrocarbons by the presence of sulfurcompounds in the charge to the intermedi-. ate zone, specified activityindexes of the variouscatalysts, the preferred catalysts, operating theintermediate and conversion zones at approximately the sametemperatures, locating these two zones in a single reaction vessel andthe preferred volumetric ratios of said intermediate and conversioncatalysts.

Other narrow embodiments of the invention rel-ate to the selected rangesof the reaction conditions within which the aforesaid control ofhydrogenation and conversion reactions is exercised. In the initialzone, these include maintaining a hydrogen partial pressure within therange of about 200-800 .pounds per square inch pressure (abbreviatedherein as p.s.i.), an hourly space velocity (ab breviated herein asLHSV) within the range .of about 0.2-15.0 based on the volume of saidliquid feed, a hydrogen charge within the range of about 500-5000standard cubic feet per barrel of said liquid feed (abbreviated hereinas s.c.f./b.) and a feed or charge temperature within the broad range ofabout 75-300". Preferred reaction conditions in the initial zone are ahydrogen partial pressure of about 300-600 p.s.i., a space velocity ofabout 0.5-8.0, a hydrogen charge rate of about 1200- t 3000 s.c.f./ b.and a feed temperature of 75250 F. For the intermediate zone, theyinclude keeping a hydrogen partial pressure within the range of about200-800 .p.s.i. (about 300-600 being preferred), an hourly spacevelocity within the range of about 2-60 (preferably about 5-40), a totalhydrogen charge within the range of about 500-1000 s.c.f./b. (thepreferred range being about 2000-5000), and an inlet temperature withinthe wide range of about 350-700 (preferably about 400-650). In saidconversion zone, the space velocity-may be between about 0.2 and 8.0,and preferably about 0.5-5.0, and the average reaction temperature isnot substantially below said inlet temperature of the intermediate zone.

In performing the instant process a feed stock with a pronouncedtendency toward undesired polymerization is subjected to a selectivehydrogenation process in which it is first hydrogenated mildly in theliquid or mixed.

phase; the resulting efiiuent is vaporized under controlled conditionsand thereafter treated in the gaseous phase with hydrogen under moresevere conditions in at least two catalytic zones. The initialhydrogenation is conducted at a temperature sufliciently low to avoid orminimize both thermal and catalytic polymerization in by drogenating asubstantial portion and usually all or almost all of the diolefins,including all of the more reactive ones, While the catalyst isrelatively fresh. During the initial hydrogenation, little, if any,desulfurization is accomplished and a substantial proportion of themonoolefins, often remains unsaturated here as, under some conditions, amore severe hydrogenation at this stage would produce substantialpolymerization.

By reason of the stabilizing effectiof the partial hydro-= genation,most of the initial efliuent can now .be vaporized without depositingpolymeric solids on the equipment,- if

the vaporization is accomplished by heatingin ,a carefully controlledmanner up until the liquid and gaseous phases are separated from=oneanother completely and rapidly in an enlar-ged separation zone; It is tobe noted that the hydrocarbon vapor phase is sharply andifcomple'telyseparated for thefirst time from :the circulating flux liquid: includingthe unvaporized -fraction of the efiiuent: by re-.

moval of the vapor phase and not by evaporating the liquid phase to,dryness or; a close approach to dryness; I This gaseous phase is furtherheated withoutanyufurther special precautions to a temperature suitablefor olefin saturation at a high reaction-rate and usuallywithin therange of 'desulfurization temperatures; then it is subjected to at leasttwo more catalytic treatments with hydrogen at distinctly highertemperatures than in the initial hydrogenation. ration of allremaining-.mono-olefins and diolefins is substantially completedandorganic sulfur compounds are converted into hydrogensulfide' Without anyappreciable polymer formation occurring, in. eitherthe preliminaryheating or in; the vapor phase reaction even though the desulfurizationcatalyst is customarily of a type of high polymerization potential. Theusual difficultieswith polymerization are not encountered in thisstagebecause of the earlier-hydrogenation of the more reactive ,diolefins.

When the catalyst in the initial reactor is inrelatively freshcondition, most of the hydrogenation of mono-ole,- fins (often more thanas well as diolefins, occurs there. So little hydrogenation of thehydrocarbons occurs in the, subsequent reactor or reactors thattheexotherms there are quite, small with little temperature differencebeing noted between the inlets and outlets. Eventually as the activityof the catalyst in the first reactor decreases with continued use, moreof thehydrogenation load is shifted .to the intermediate hydrogenationcatalyst bed in the second reactor and substantial increases between theinlet and outlet temperatures of this reactor are then apparent. i

This shift of the hydrogenation load demonstrates im-:

portant advantages of the novel process In the absence of anintermediate catalyst of high hydrogenation activity operating at atemperature higher than the initial zone, appreciable proportions offairly reactive unsaturates, including some diolefins, would-reach the,desulfurization catalyst much earlier. than in :the present processand: deactivate this catalyst sooner by depositing, solids there-Accordingly, the present invention enables the de on. sulfurizationcatalyst to remain on-stream i for longer periods before regeneration isrequired to restore its cat-- into an additional stage of eithertheinitialhydrogena-i tion catalyst bed or the final desulfurizing ;bed ofany known two-stage hydrogen treatment. Theaverage reaction temperatureis substantially higher, usually at least about 200-500 higher, in theintermediatezone than in the initial hydrogenation zone and a higherheat level.

substantially increases the hydrogenating activity of the catalyst;hence, some .unsaturated aliphatic hydrocarbons of less reactive natureare saturated'in-the intermediate zonewhich would not react .withhydrogen to any sub-- stantial extent at the lower temperature in afirst zone of extended size. Also, the intermediate catalyst has a dis?In the final conversion, the satu-;

tinctly higher hydrogenation potential at any given temperature than thedesulfuriza-tion catalyst, so increasing the size of the desulfurizationcatalyst bed would not produce comparable results; and this is true evenwhen the intermediate catalyst and the desulfurization agent are locatedin beds adjacent to one another in the reactor and operating at aboutthe same average temperature.

While it is preferred to locate both the intermediate hydrogenationcatalyst bed and the final conversion or desulfuriza-tion catalyst inthe same reactor for purposes of simplicity and economy, this is notessential and the two catalysts may be located in two or more differentreactors. Whether located in the same or separate reactors, the two bedsmay be operated at about the same average reaction temperature. Little,if anything, appears tobe grained by either heating or cooling thegaseous stream between these two stages. Exothermic heat is generated inthe final conversion catalyst, usually to a small extent, so that theaverage temperature at this stage may be slightly above that of theeffluent of the intermediate hydrogenation zone. However, in some casesthe loss of heat through the reactor walls may exceed the exotherrn andresult in a final catalyst bed slightly cooler than the intermediatehydrogenation catalyst bed. The ratios of catalyst volume in theintermediate bed to catalyst volume in the final conversion bedgenerally range between about 1:40 and 1:2.

The arithmetic average reaction temperature in the initial hydrogenationreactor usually runs about 20 to 80 above the temperature of the chargeentering the reactor. In the intermediate hydrogenation zone, thereaction temperature is usually found to average between about and 70above the temperature at its inlet. The average reaction temperature ofthe final conversion or desulfurization catalyst zone is about the sameas the temperature of the charge entering there. To mention a fewtypical figures, the average temperature of the initial hydrogenationbed may be about 125 F. for palladium, about 200 for platinum and about230 for a nickel catalyst; and for the intermediate hydrogenation zone,the average temperature for the same three catalysts may be about 375,400 and 425, respectively. When the desulfurization catalyst comprisesthe oxides or sulfides on cobalt and molybdenum on an alumina support,the average temperature in the final catalytic zone may be about 350-700and preferably about 400-650".

As the starting material, any mixture of aromatic and unsaturatedaliphatic hydrocarbons with organic sulfur compounds may be employed inthe present process if the final boiling point of the liquid does notexceed about 500. A narrow boiling range material, for example, onehaving a boiling range between about 140 and 275, is desirable andpreferably a charging stock boiling in the range of about 160 to 220 forproducing benzene.

As a source of aromatic hydrocarbons, the liquid feed desirably containsa total of between about 20 and 90% aromatics, especially benzene andtoluene. Typically, it also has substantial contents of diolefins andolefins as evidenced by Diene Numbers of about 10 to 25, which measurethe proportion of conjugated diolefin as determined by the maleicanhydride condensation method and Bromine Numbers of about to 30, whichrepresent the total content of unsaturated aliphatic hydrocarbons. Feedswith Diene and Bromine Numbers somewhat higher than about 45 and about75 respectively may also be processed according to the presentinvention. The organic sulfur content is typically about to 300 ppm. andmay be as high as about 1000 ppm.

In other utilizations of the present process, the charging stock neednot be rich in aromatic hydrocarbons. For instance, in producing astable gasoline blending stock from a pyrolysis liquid, a feedcontaining 6 to 20% aromatic compounds is typical.

Feed stocks of the nature described are unstable as they tend to formpolymeric gums readily. It has been found desirable to keep the periodof storing them as brief as possible in order to minimize theintroduction of gum into the present process. In addition it isrecommended that the liquid feed stock be free of dissolved oxygen andbe stored in the substantial absence of oxygen or air, for example,under a blanket of an inert gas such as nitrogen. This prolongs theactivity of the catalysts usable in this process. Such feed stocks aregenerally obtainable by severely thermally cracking a petroleum fractionsuitable for the manufacture of gasoline or light olefins as exemplifiedby ethylene. It is preferred here to depentanize the cracked product. Aparticularly preferred feed is one with an end point not exceeding 220and a maximum gum content of less than 15 milligrams per milliliters.

The total consumption of hydrogen in this process varies of course withthe particular feed stock employed, but in general, it is in the rangeof about -900 s.c.f./b. of liquid feed stock. A typical value is 300s.c.f./b. with a charging stock of Diene and Bromine Numbers of 15 and24 respectively. The consumption is usually found to be less than 500s.c.f./b. Substantial excesses of hydrogen have been specifiedhereinbefore to maintain a sufficient hydrogen partial pressure to avoida drop in the degree of hydrogenation. Although pure hydrogen may beused, it is customarily supplied as a mixture of hydrogen and gaseoushydrocarbons in the off gases of units for reforming naphthas orhydrodesulfurizing gas oils, etc. The gas charge preferably has ahydrogen content of at least 60% but gaseous mixtures with as little as40% hydrogen may be used, such contents referring to percent by volumeor molar percent.

The partial pressure of hydrogen in the two or more reactors isimportant in avoiding undesired side reactions, such as the formation ofgum or coke on the catalysts. It should be maintained within the rangeof about 200-800 p.s.i., the 300-600 p.s.i. range being preferred. Thetotal pressure in the reactors is not critical but it should not be sohigh as to interfere significantly with the vaporization of the feed andreaction products described herein. Typically, a major proportion of theproduct gases with much unconsumed hydrogen is recycled to the processafter any excessive quantities of hydrogen sulfide have been scrubbedout and this usually constitutes a major proportion of the totalquantity of gases charged to the reactors.

The charging stream of combined recycle and make-up gases containinghydrogen is divided into several streams. A substantial quantity ofhydrogen must be introduced into the first reaction zone along with theliquid feed, and unreacted hydrogen is present in the eflluent of thatreaction which is subjected to further hydrogenation reactions. While intheory all of the hydrogen-rich gas required for the series of selectivehydrogenations can be charged to the initial reactor, this is notparticularly desirable in practice. Especially since the circulating gasmay be employed after heating to a high temperature in a furnace as aheat source to aid in vaporizing the efiluent from the initial reactorand also to regulate the temperature of the vapor phase charge ofhydrocarbons and hydrogen to a subsequent reactor. Alone thishydrogen-rich gas stream can be heated without decomposition or otherdifficulty to a temperature several hundred degrees higher than it ispossible to heat the liquid or mixed phase effluent of the first reactorwithout the coincident deposition of polymeric gum or coke. Suchdeposition from the mixed phases can occur at temperatures of 300 oreven lower. In serving as a heat source, a substantial part of the totalcirculating gas, say about 30 to 85%, is heated to a temperature in therange of about 500 to 950, and preferably in the range of about 600850,while the unheated balance of the gas is charged to the initial reactor.One stream typically containing more than half of this heated gas isused to supply the final temperature increment to the initial mixedphase eflluent just prior to entering the enlarged separating andvaporizing chamber and the remainder may be introduced into the whollygaseous streani leaving the top of said chamber on its way to the secondstage reactor as the final heat increment to adjust the charge to thedesired inlet temperature.

A catalyst of high hydrogenation activity is required for the initialreaction zone as it must hydrogenate at a relatively low temperature themore reactive conjugated diolefins and usually at least some of theother'olefins, but the polymerization activity of this catalyst must berelatively low in order to avoid the formation of gums which willdeactivate it. While some suitable hydrogenation catalysts alsoincidentally possess relatively high desulfurization activity initially,this property usually'drops oif rapidly in a period of a few days to aweek, because such catalysts are readily poisoned with respect todesulfurizing ability at desulfurizing temperatures by feeds containingmuch organic sulfur.

These qualities of the catalysts may be defined in terms of arbitraryactivity indexes which are described herein. Unless otherwise stated,all such indexes are measured using fresh new catalyst. The activityindexes enable one to clearly differentiate between the two or morecatalysts employed at various stages in the instant process.

For delineating hydrogenating activity, two different indexes areavailable. is defined herein as the percentage or proportion of isoprenewhich is converted to pentenes and pentanes when a blend of 8-10%isoprene and 50-500 ppm. of thiophene sulfur in benzene is passed overthe catalyst with 1500-3000 7 s.c.f./'b. of hydrogen gas at 150 F., 300.pounds per square inch gage (hereinafter designated p.s.i.g.) as thetotal pressure and a liquid hourly space velocity of 5. Thus aconversion of half of the isoprene present, or 4.5% out of a total of9.0% isoprene present, signifies that the activity index is 50. For theinitial and intermediate zone catalysts, a hydrogenation activity indexof at least about 40 is recommended.

In determining the benzene conversion index as another and usuallysupplemental measure of hydrogenation activity, a sulfur-free mixture of17% benzene and 83% cyclohexane is passed through the catalyst under;

test at 400 F. and 400 p.s.i.g. with 1500-3000 s.c.f./b.

The hydrogenation activity index hydrogen circulation and a liquidhourly volumetric space 'benzene conversion index of at least about 50,meaning j that half of the benzene present or 8.5% is converted intocyclohexane, but an index of about 100 is typical for the preferredcatalysts.

Another means for designating suitable catalysts for the first andintermediate zones is the polymerization activity index. This is anotherarbitrary index which equals the percentage of isoprene that ispolymerized when 25' cc. of a mixture of 8-10% isoprene in benzene isheated with 5 cc. of the catalyst to be tested in a stationary'bomb of30-55 cc. capacity to a temperature of 350 under a blanket of inertnitrogen gas and held there for one hour. The polymer formed from theisoprene remains on the catalyst, or the interior surface of the bomband the liquidconsisting of benzene and unreacted isoprene is poured offand analyzed chromatographically. The decrease in isoprene monomer duetopolymerization is calculated by difference between the isoprenecontent of the reaction product and that of the test blend charged. Asatisfactory catalyst for the initial and intermediate stages has apolymerization activity index less than about 35, as polymerizationthere is undesirable. 7

An arbitrary desulfurization activity index may also be used in thepresent invention, principally for determiningsuitable catalysts for thesubsequent desulfurization operation: This index .is 'theperceritre'duction in sulfur content obtained when a blend of purecompounds.

consisting 'of 10% hexene and 10% isoprene. in vol- 111116, percent ofbenzene with 'a total thiophene sulfur; 7

content of 500 ppm. is passed over the catalyst in question at 500 F.and: 450 p.s.i.g. together-with between 1500 and 4000 s.c.f./b.ofhydrogen at a liquid hourly volu-.- metric space velocity of 2.. Forsuitable, desulfurization,

the finalstage catalyst desirably has a desulfurization operation theactivity index is in the 0 to 50 range.

Catalysts of substantial acid activity are usually not desirable' forthe initialgiand 'intermediate stages of this process since theyproduceunwanted cracking reactions, sov silica-alumina catalyst supports areusually avoided.

However, while it is preferable that the catalyst support besubstantially free of halogens ,,a relatively low halogen;

content up to about 0.5% may be tolerated; Further: more, a catalyst isfavored which is substantially devoid of alkylation activity and thusdoes. not promote the alkylation of aromatics with olefins.

Although they are not necessary, porous particle form supports arerecommended for all of the catalysts used in the instant inventiont-oadequately disperse and increase .the surface area of the actualcatalytic agent. In

the case of expensive agents such'as platinum,'the support'is quiteimportant from an economic. standpoint in commercial-scaleoperations;The particle form support may take the physical form of pellets, rods,tables, spheres or granules of irregular shape. The average particlesize may range from inch to /2 inch.

The porous supports may take the form of natural or treated clays, suchas Fullers earth, kaolin, bentonite,

montmorillonite and superfiltrol; treated clay-like materials, such.-ascelite and sil-o-cel; artifically prepared or synthetic materials, suchas magnesium oxide,.silica gel alumina gel, and the like; or thezeolites, activated carbon, di-

atomaceous earth, kieselguhr, infusorialearth and the like.- Adsorptivevaluminas, bauxite or ;porocel are particularly desirable carriermaterials. .Activated alumina, a well-known crystalline alpha. aluminamonohydrate prepared by partial dehydration of crystalline alpha aluminatrihydrate, is very satisfactory for the purpose. Another highlysatisfactory form of alumina is currently marketed by the AluminumCorporation of Amer.- ica under the name F-IO Alumina. F-10 Alumina ischi alumina vprepared by the precipitationzof alpha alumina trihydratefrom an aluminate solution below F. After filtering and drying, thismaterial is' calcined first between 536 and 842? Fito effect. partialdehydration, and later at 750 to .1470-.

A varietyof catalyst of ditfering chemical constitu-j tion may beemployed :in the-initial and intermediate hydrogenation steps as long asthey have the necessary Platinum in amounts rangactivity describedherein. ing from about 0.05 to 2.0%, preferably about 0.2-1.0%,

supported ion various aluminas and especially. gamma and chi alumina, issuitable .as are the other noble materials in Group VIII of the PerodicTable of Elements with atomic numbers of at least 27; such as rhodiumand palladium. The concentration of palladium in such catalysts may 'be.about 0.0510% and about 02-20% is preferred for the purpose... Withinthe latter limits,

the hydrogenation activity of thelcatalyst increases whenthe palladiumcontent is increased. Nickel,:either unsupported or on unknownsupporting materials in concentrations ranging down to about 10% nickelin the composite catalst also provides satisfactoryresults, as

does copper chromite. For instance, good hydrogenatingcharacterististics for the first reactor are obtained with 55% nickelsupported on kieselguhr-a composite that is strongly selective inhydrogenating diolefins in preference to olefins. By reason of theirhigh hydrogenation activity at low temperature, palladium or platinum ongamma alumina are recommended, palladium being preferred for the initialreaction, since its greater activity catalyzes the desired hydrogenationreactions at a temperature about 100 lower than in the case of platinum.The palladium composite is desirably promoted in some instances with aquantity of chromia in the same range as the palladium. Platinum on asupport of alumina or another suitable material is preferred for theintermediate reaction stage because of its apparently superiorresistance to the higher temperatures in that zone. Among the manysuitable specific catalysts are palladium on activated canbon and 0.6%platinum on eta or chi alumina of less than 0.01% chlorine content. Themanufacture of such catalysts is well known in the art and accordinglyis not described here.

The catalysts employed in the second and any subsequent reactor operateunder quite different reaction conditions from those in the initialreactor. The charge is entirely in the gaseous phase and thetemperatures employed are substantially higher than in the initialhydrogenation zone. As a result of the substantially greater averagetemperature in the intermediate catalyst bed producing enhancedhydrogenation activity, hydrogenation proceeds there with furthersaturation of diolefins and usually of a substantial proportion ofmonoolefins also, even where the feed stock has been partially saturatedby a treatment with the same catalyst in the first reaction Zone. In thefinal stage, the reaction products leaving the intermediate catalyst bedare further treated with hydrogen in the presence of a desulfurizationcatalyst to desulfurize the charge and to substantially complete thehydrogenation of the less reactive unsaturated hydrocarbons therein,namely the monoolefins and any remaining diolefins. The desulfurizingreaction requires a temperature sufficiently high for organic sulfurcompounds to be converted into hydrogen sulfide at a commerciallyacceptable rate or space velocity. With the catalyst commonly employed,this means that the inlet or charge temperature should be at least about350, and preferably at least about 400.

The conversion catalyst in the last reaction stage may be any knowndesulfurization catalyst including tungsten disulfide, nickel sulfide,nickel-tungsten sulfide and chromia on alumina promoted by molybdena.Comlbinations of a compound of a metal in Group VI B of the PeriodicTable of Elements, for example a chromium, molybdenum or tungstencompound, together with a compound of a metal of the iron group providegood results and it is usually desirable to support such agents on aconventional catalyst support, such as alumina, kieselguhr, etc. Theoxides or sulfides, particularly the latter, of cobalt and moylbdenumsupported on gamma alumina are preferred for the desulfurizingoperation. In catalysts of this type the concentration of cobalt maynange from about 1 to 5% while that of molybdenum may be 'bewteen about5 and 18%.

The final catalyst may also be defined in terms of an arbitrarydesulfurization activity index as set forth hereinbefore in the 80100range which activity is retained for a period of at least one week andusually much longer. It should also have at least moderate hydrogenationactivity. The increased temperature of the final reaction greatlyincreases the actual hydrogenation activity of these catalysts. Also ithas been found that some and perhaps all catalysts which are well suitedfor the final desulfurization reaction have relatively highpolymerization activity indexes exceeding about 25 on the scale definedhereinbefore, even though such activity is t0 neither directly concernedwith nor desired in the instant process. By reason of the prior partialor complete hydrogenation of the more reactive diolefins in the chargeto the final stage, there is little tendency for polymerization tooccur.

In illustration, the presulfiding or final step in the preparation of apreferred type of desulfurization catalyst comprising sulfides of cobaltand molybdenum on alumina may desirably be performed in situ in thefinal reactor, provided that any noble metal or other catalystsusceptible to poisoning by hydrogen sulfide is absent from the reactor. Otherwise either the noble metal catalyst should be removedbefore this operation, or the sulfiding step should be performedelsewhere or the reactor should be purged with a hydrogen-rich gas forat least 10 hours after sulfiding, and preferably about 20 hours, atnormal average reaction temperatures. In this illustration, a freshcontact catalyst containing cobalt molybdate on the surface of asuitable support such as gamma alumina or a catalyst regenerated to theoxide state by combustion with air diluted by steam is subjected firstto pre-reduction for six hours at 700 p.s.i.g. and 700 with ahydrogenrich recycle gas substantially free of hydrogen sulfide.Following the prereduction step the catalyst is then contacted with acirculating stream of mixed hydrogen sulfide and hydrogen underconditions such that the minimum partial pressures are 8 p.s.i. forhydrogen sulfide and 100 p.s.i. for hydrogen and the temperature is inthe range of 500 to 700. The treatment is concluded at a temperature ofabout 700 after being continued until the sulfur content of thecomposite catalyst rises to the range 6.5- 7.5% whereupon the catalystis ready to be placed onstream. Later during desulfurizing operationsunder a hydrogen partial pressure of 300 p.s.i. or more, the sulfur inthe catalyst drops from that range to an equilibrium content of about4.6%.

Returning now to the first reactor, suitable ranges of reactionconditions have been described earlier and the actual reactionconditions are selected and regulated within those ranges in a mannerknown to those skilled in the art to produce an initial hydrogenationefiluent in which desirably at least about 35%, and preferably at least50%, of the original diolefins have beeen converted into mono-olefins orparatfins and in which an amount equal to at least about 25 andpreferably at least about of the liquid feed rate remains in the liquidphase. It is also preferred to obtain an effluent liquid fraction with aBromine Number at least about 25% below that of the liquid feed.

The regulation of such reaction conditions and the effect of oneoperating variable upon another are well understood by those skilled inthe art and need not be explained in detail here. For instance, if thedegree of hydrogenation tends to drop below the minimum specified, orperhaps below the preferred values, this condition can be corrected byincreasing the feed temperature or decreasing the space velocity orboth. Also if the proportion of initial reactor effiuent in the liquidphase drops below the minimum specified, the feed temperature may bedecreased, the space velocity increased to reduce the total exothermicheat generated and provide a greater quantity of reactants to absorb theheat liberated, or the pressure increased or any combination of thesemeasures may be employed in reducing the degree of vaporization in theinitial reactor. Using the same circulating gas, an increase in totalpressures of course results in a corresponding increase in hydrogenpartial pressure.

With a fresh catalyst, either new or regenerated, it is obviously mosteconomical to maintain the feed or charging temperature at the lowesttemperature at which the gaseous and liquid components of the charge arereadily available thus avoiding any heating or cooling expense andminimizing gum formation. The feed temperature is desirably within thestated range and, in the preferred operation, the feed temperature ismaintained at a substantially constant value within the narrow range of75- it 190 F., in the lower part of that range being recommended, whilethe catalyst is fresh. Usually, this tem- I perature is subsequentlyincreased either gradually or by steps but not beyond about 300 F. inorder to maintain a diolefin reduction of at least 35%, and preferablyto maintain a substantially constant degree of saturation, as I is lessthan 35%, or the degree of vaporizationexceeds 80%, or both, after thefeed temperature has been adjusted upward to the maximum temperaturethat is conr sidered suitable for the particular catalyst. These arebetter criteria than prescribing a maximum outlet temperature for theinitial reactor inasmuch as the degree of vaporization of the effluentand the degree of saturation of its more reactive original components,are more significant than the outlet temperature in the instant process.In addition, it appears that the maximum permissible outlet temperaturecan vary considerably for different feed stocks over therange of about275 to 400. instance, a reactor outlet temperature of 325 is consideredexcessive for certain low boiling feed stocks but will give satisfactoryresults with other. feeds boiling at higher temperatures ranging up toend points near 500 It is further apparent that an attempt to defineacceptable outlet temperatures is not feasible in consideration of thevariations in permissible pressures in the reactor. For example, anoutlet temperature of 325 would be suitable for a relatively highreaction pressure in retaining the desired proportion of effluent in theliquid phase while a lower temperature would be necessary if the minimumpressure were employed while all other con ditions were held constant.Thus, as a result of the close interrelation of the various operatingconditions, it is For more significant to described a preferredembodiment of the initial hydrogenation in terms of the regulation ofcertain reaction conditions within restricted ranges to provide anintermediate product in which a certain proportion is retained in theliquid phase and a certain amount of the more reactive feed componentsare at least partially saturated.

An entirely ditfe'rent situation prevails at the outlet of thedesulfurization reaction zone as it is unlikely that any exothermcreated by the reaction conditions mentioned herein can reach atemperature sufficiently high to deactivate the catalyst. However, topermit the use of ordinary construction materials, the maximum outlettemperature should not exceed about 850".

Although catalysts in the form of palladium or platinum supported onalumina retain sufficient activity for extremely long periods, as forinstance, 5 months and usually more in the case of palladium catalysts,regeneration of the catalyst is eventually necessary. This may bereadily accomplished by heating the reactor to a temperature of about 700900 for a palladium-alumina bed While passing a gas containing 1 to 2%oxygen therethrough. A diluent is usually introduced with the air toavoid excessive regenerationtemperatures which can reduce catalystactivity considerably. Nitrogen or flue gas may be used generally forthat purpose and the more convenient medium of steam may be utilized asthe diluent with a palladium catalyst. The regeneration of othertionally regenerated in similar fashion at even longer" intervals ofabout 9 months or more when the organic sulfur" content of the finalreactor-efiiuent: exceeds 20 p.p.m. even when the inlet temperature ofthat reactor.

has been raised to the permissible maximum. This regeneration convertsmOstcf-the cobalt and molybdenum compounds to oxides and apresulfidingtreatment'such as the one descirbe'd hereinbefore ispreferably employed to restore the catalyst to its original more activesulfide form.

It has also been-found that purging the initial and intermediate contactmasses withhydrogen at 200500 p.s.i.a. and 750850 for 16-4 hourssometimes serves to regenerate certain catalysts, suchas palladium,almost as effectively as conventional regeneration by combustion withair diluted to an oxygen contentof 1 or 2 percent. Accordingly, it is.contemplated that, in the ab sence of sever deactivationofftheicatalyst, these catalysts may be regenerated 'severalfltimes bysuch treatmentwithhydrogen-rich gas before ,it is necessary toregenerate them by the combustion technique.

Only a limitedamount of hydrogen-sulfide may be,

tion stages. Althoughthis loss of activity, may be readily restoredeither by regeneration 'of the catalyst in the usual fashion orthe hothydrogentreatment described earlier, frequent regenerations reduce the.over-all efii ciency .of the process. Accordingly, in the. ;case.of aplatinum catalyst supported on alumina, it is desirable that the partialpressure of hydrogen sulfide in the gaseous phase should notexceed 0.05psi. and preferably should be less than 0.03 p.s.i. The effect on apalladium catalyst is similar.; Organic sulfur generally has a lessereffect on the catalyst and it is'a relatively simple matter. to controlthe hydrogen sulfide: which is introducedin the hydrogen-containing gasby simply passing either; or both of the make-up andrecycle gasesthrough an 'alkaline scrubber, or other unit for removing hydrogensulfide.

such as a diethylamine absorber.

Under severe conversion conditions, for example a high desulfurizatio'ntemperature in combination with a 10Wv space velocity of perhaps lessthan 1,; a sulfided composite of cobalt and molybdenum on alumina maycatalyze the hydrogenation of a part of. the aromatic: hydrocarbons, asexemplified by the conversion ofbenzene. to cyclohexane.

This is usually undesirable and maybe .easily avoided by similarlysevere conversion conditions, and the same in Where hibiting treatmentis: useful, in this. zone also. the charge. contains less of suchsulfur, compounds it is a simple matter to supply additional hydrogensulfidein the hydrogen-rich gasv which is introduced upstream of thefinal reactorx Selection of a make-up gas of. suitable hydrogen sulfidecontent or by-passing the recycle gas around the caustic soda scrubberare some of the methods useful in attaining any additional inhibitingeffect.

Despite the unstable nature of the hydrocarbon feed stock, very littleif any gum is formed in the firstreactor.

The relatively low' reaction temperature is not conducive to thermalpolymerization. A cataly'sthaving little or, no polymerization activityis employed. (A substantial pro-f portion ofthe reaction mixtureismaintained .in the liquid phase to avoid approaching the pointof'drynessin the; reactor. In addition, the usually substantial aromaticC011,

tent of this liquid makes it a good solventtor-polymeric. gums, so theliquid phase flowing downward-1y through this mixed phasereactordissolvesand carriesgalong insolution most of any polymer formedtherein. 1

The; second catalytic reaction with hydrogen is entirely a vapor phaseoperation; hence,,it is necessary to vaporize most of the diluent of thefirst reactor. Accomplishing this by merely passing the initial efiluentthrough a heater and into a second reactor is not satisfactory eventhough this technique has been suggested in the prior art. Suchprocedure deposits polymer either in the heater or in the next catalystmass or both, and stoppages of this nature call for much cleaning and/or regeneration that reduce the overall operating efficiency.Accordingly, vaporization of the initial efiluent in the presence of aflux liquid is preferably employed here. This may be accomplished byvarious methods, one of which involves a combination of stages in whichthe initial hydrogenation efliuent is gradually heated under goodtemperature control in the presence of a flux, preferably circulating insubstantial quantity through the transfer line between the initialreactor and a vaporizing and separating chamber of enlarged crosssection. In that chamber vaporization of the initial original feed andproducts thereof is completed to the desired extent of about 90 to 99%and seldom more than 99.5%. The small but significant balance ofunvaporized efiluent is withdrawn at least intermittently from theprocess as a liquid leaving the enlarged chamber and it carries a smallamount of polymer formed during the vaporization operation and possiblyalso in the initial hydrogenation step or perhaps present in theoriginal charge stock. Once this separation of the gaseous and liquidphases has been accomplished, there is no longer a tendency toward anysignificant polymerization in the gaseous phase containing the majorproportion of the hydrocarbons even when it is heated up to temperaturesof 350 to 700 which would have produced an unacceptable degree ofpolymerization in the mixed phase material from the initial reactor.

The gradual heating of the initial efiiuent to effect controlledvaporization during passage of the efiiuent through the restrictedtrans-fer conduit (including heater passages, etc.) leading from theinitial reactor to the vaporizing and separating chamber may beaccomplished by several means. One, comprises an optional but preferredtechnique in which a circulating liquid flux at a substantially highertemperature, than the efiluent typically of the order of 75-200 higher,is injected into the initial hydrogenation effluent near the outlet ofthe first reactor. It will be appreciated that the exortherm of theinitial reaction has already increased the temperature of this effluentsubstantially above the temperature of the feed to that reactor. Thetemperature of the mixture of flux and reaction eflluent is preferablyincreased further during passage through an indirect heater which isdesirably heated with steam or another easily controllable medium foreven heating. A relatively low temperature difference between theheating and the heated media is highly desirable to provide the gentleheating that minimizes polymerization in such equipment. Indirect heatexchange is recommended for the major heat input into the stream passingthrough the transfer conduit. Finally, and preferably closely adjacentto the inlet of the separator pot,

an additional stream of the hydrogen-rich gas used in this process maybe injected into the mixture at a substantially higher temperature up toseveral hundred degrees higher than the temperature of the mixture. Thisdirect contact heating with jet of hot gases is an optional but highlydesirable feature which minimizes polymer deposition on equipmentsurfaces. With each of these increments of heat, more of the firstreactor efiduent is converted in the transfer conduit from the liquidphase into the gaseous state under conditions in which the presence atall times of a substantial liquid phase assists in preventing or atleast in minimizing the deposition of polymeric material on heatedsurfaces. The enlarged cross-section of the chamber provides goodconditions for separating the two phases by reducing the vapor velocitysufiiciently so that all of the liquid drops out of the rising gaseousphase.

The supply of steam to the indirect heater may be manually controlled tomaintain a predetermined temperature in the separating chamber as steadyas possible, but better results are usually obtainable in regulating thesteam supply in response to the liquid level in the separating chamber.That regulating system involves controlling the input of steam manually,but preferably automatically, in direct response to the signals of aconventional liquid level indicator or controller attached to thevaporizing and separating chamber. The removal of liquid streams fromthat chamber as Well as any input of external flux is desirablymaintained at constant flow rates under the regulation of automatic flowcontrollers; therefore, a rise in the liquid level in the separatingchamber represents a decrease in the vaporization of the initialhydrogenation efiluent and a fall in that level means that the effluentis being vaporized in a greater degree. To maintain a steady degree ofvaporization more steam or less steam respectively is supplied to theindirect heater. The heating steam may be adjusted by means of a valvein the steam supply line or one in the line used for draining condensedheating steam from the heater.

Conventionally, control of vaporization of a generally similar nature isregulated in response to the temperature of the vapor or perhaps theliquid temperature. Such control is subject to the usual deviationsencountered in efforts to obtain precise elevated temperaturemeasurements that arise from radiation or evaporation of liquid on atemperature sensing element, etc. Moreover, it is not particularlysatisfactory for liquids of narrow boiling range, such as the preferredfeeds of the present invention, inasmuch as a small temperaturedifferential of only a few degrees at a substantially elevatedtemperature generally is related to a large differential in theproportion of liquid vaporized. Thus control of heating of the liquid indirect response to the actual proportion of unreacted feed stock plusreaction products (initial effiuent) retained in the liquid phase ismuch simpler and far more accurate here than control based on theindirect factor of temperature which is further influenced by variationsin pressure, in the composition of the liquid, in hydrogen to liquidfeed ratios and system lag.

Either manual or automatic control of the heating of the initialeffluent in direct response to the liquid level in the flash chamber mayalso be extended to controlling the quantity of heat supplied by thestream of hot hydrogen-rich gas injected into the transfer line near theinlet of the separator pot. This regulation may govern either thequantity of said gas being admitted to the transfer line or thetemperature at the charge outlet of the furnace described hereinafterfor heating that gas. Also, it is possible to control both the heatinput to the indirect heater through which the initial efiluent passesand the heat furnished to the effluent by the hot hydrogen-rich streamin response to the liquid level controller on the vaporizing andseparating chamber. However, it is usually preferred from a standpointof practical operations to apply such regulation only to the steam inputto the indirect heater.

The flux liquid comprising the liquid fraction of the effluent from theinitial reactor and any inert liquid miscible therewith that isintroduced into the transfer line may perform several functions beforebeing separated from the gaseous portion of that efiiuent in theseparation chamber. It minimizes or inhibits gum formation at thiscritical stage of the preferred process wherein a stream of mixedgaseous and liquid hydrocarbons containing gum-forming precursors iscarried to a relatively high degree of vaporization by heating, for theflux prevents the efiluent from approaching dryness too closely, forexample, not closer than about 5% based on the original liquid feedrate. Secondly, the circulating flux serves as an economical andrelatively gentle direct heating medium for vaporizing a portion of theinitial effluent. Finally, the flux liquid prevents, or at leastminimizes the deposition of any gums or polymeric solids on the pipesand other apparatus by reason of its washing action on the surfacesthereof and its solvent characteristics which enable it to retain insolution any polymeric material whether formed at this stage or earlier.

Although any hydrocarbon liquid of suitable boiling and stabilitycharacteristics may be employed as the fluX,

it is preferred that the content of aromatic compounds should amount toat least 15% to improve its capability for dissolving gummy material- Aflux liquid from an external source may be used, and it is suggestedthat its volatility should be sufficiently low that a major propor-.

tion and preferably substantially all of the flux remains in the liquidstate under the conditions in the vaporizing chamber while itsresistance to coking and polymerization should desirably be at least asgood as that of the initial eflluent. Its boiling range is preferablylocated between about the boiling point of benzene and about 950.

However, an economical and readily available flux liquid may be obtainedvery simply by merely reduc-.

amounts to accmulating the least volatile fraction of the 5 feed stockas the liquid flux.

The rate of recirculating the flux liquid may amount to at least 5%, andpreferably at least of the rate of introducing the liquid feed stockinto the first reactor, and lesser amounts may be recirculated where anappreciable proportion of the initial efiluent is retained in the liquidphase throughout the vaporizing step. As used herein, all flux (liquideflluent plus any added liquid) quantities or dates relate to theproportions at the moment when the maximum degree of vaporization of theinitial effluent is attained; and, of course, the proportion of materialin the liquid phase reaches-its minim-umnamely the instant of separationof the gaseous and liquid phasesrather than at the confluence of acirculating flux stream with the initial hydrogenation efiluent. Muchhigher flux circulating rates can be employed ranging up to and even to200% or more, for the only real limitations are physical ones relatingto the capacities of the equipment and economic ones relating to pumpingcosts and the cost of larger equipment. When the total proportion ofliquid in the. transfer line and heater is ample by a substantial marginto avoid dryness and bathe the walls of the equipment, further increasesin the flux circulation vrate do not achieve a corresponding or even asignificant reduction in the amountof polymer formed in the system oreven the efiluent of the initial reactor or from a supply of externalflux or from both sources, and over any substantial period the rate ofwithdrawal must equal the supply from these sources. Under the preferredsteady state conditions, reducing the degree of vaporization of theinitial effluent and correspondingly increasing the spent fluxwithdrawal results in a decrease in the polymer concentration in thecirculating flux and vice-versa. As indicated earlier, this removalofspent flux liquid amounts to at least about 0.5%, and desirably aboutlto 10%, based on the liquid feed rate. While the amount may be larger,it is generally uneconomical to withdraw much more in the liquid phasefor purification or further processing. In actual practice a flowcontroller on the spent flux line from the vaporizing chamber may beadjusted manually as needed to keep the gum content of the circulatingliquid low enough to avoid the deposition. f polymeric material in theequipment; for example,

16 by keeping the gum content below about 200 milligram per 100milliliters.

Where a flux liquid from an external source is supplied to the system ata constant and usually relatively low rate, it is. possible to vaporizea correspondingly greater'proportion, in fact the whole of the liquidefilu- However, it is preferable to retain ;the least volatile 0.5 or 1%of said eflluent in the liquid state'in order to keep the temperature,as low as possible during, the vaporizing operation. For

ent fraction of'the firstreactor.

example,-with all percentages based on the liquid feed rate, one maycontinually charge, 5% .of a hydrocarbon oil havingan atmosphericboiling ,rangeof 600-700 and a major proportion of aromatichydrocarbonsto the separating: chamber as circulating flux, and recycle 25% liquidfrom :this pot to the transfer line immediately downstream ofthe firstreactor; then :vaporization ofithe efiiuent-flux mixture in the transferlinermay. be controlled by appropriate heating to' retain 1% of: theinitial spent flux may be withdrawn continually from the bottom thereofin maintaining steadyoperations.

The size and shape 'of'the separating and vaporiz- In avoiding orminimizing ing chamber are not critical. appreciable entrainmentof'liquid droplets in the vaporous phase that .is leaving, it is desirableto keep .the ve-. locity of the gaseous phase relatively low, perhaps 2feet. per second or less. a reasonably large cross-sectional areaperpendicular :to. the direction of gas flow in the upper part of the,vessel.

On the other hand, where the heat. for vaporization is regulated inresponse to liquid level in'the chambenit is desirable to have arelatively small across-sectional area in the. neighborhood of thatlevel in order. that a signifi cant change in level, will occur whenevera significant change in the degree of vaporizationof initial efiluentoccurs. is no necessity for maintaining aconstant cross-sectional areathroughout'the length of the chamber. As. one illustration, the,vesselmay be in thegeneral form of a double cylinder having .a lowersection of considerably.

smaller diameter than the upper section.

After separation of. the flux liquid from gaseous material derived fromthe eflluent of the initial reactor,-

this gaseous phase isheatedto bring its temperature up to the desiredinlet temperature of the second reactor and its proportion ofhydrogen-is boosted, if necessary, to the desired level for that reactorby;the introduction of a hydrogen-rich gas. These steps may be combined,if so desired, by introducing the extra hydrogen-contain-; ing gas at asubstantially greater temperature, sail about to 400 'more,.than that ofthe gaseous phase leaving the separating chamber.: This is one of thesuitable methods of making the final temperature adjustment in thecharge to the .secondreactorp It is preferably accomplished byregulating the volume of fuel gas burn-. ing in a furnace for heatingcirculating gas and consequently the outlet temperature of {thatcirculating gas streameither manually or automatically in response 'tosignals from. a temperature sensing device located in the conduitleading to. the inlet of the second reactor.

For a better understanding of the nature and objects of this invention,reference should be had to the detailed description and exampleshereinafter taken in conjunction with the accompanying drawing whichisia simplified fiow' sheet or schematic representation-of the processof the:

especially instrumentsfor indicating. recording, or 'Iegu- Y latingtemperature, pressure, level, flow, etc.

This can be achievedby providing Such factors pose nogreat problems, asthere 17 EXAMPLE 1 Turning now to the drawing, a freshly-distilledstream of thermally cracked and depentanized gasoline (160220 B. R.) ofthe composition set forth in part in column 1 of Table I hereinafterenters the feed conduit 2 at ambient temperature and a pressure of 740p.s.i.g. at the rate of 2300 b./d. This narrow cut is substantially freefrom two easily polymerizable and therefore particularly troublesomecompounds, namely cyclopentadiene which boils about 106 and styrenewhich boils around 293. A fresh or make-up gas rich in hydrogen andobtained from the off-gas of a conventional unit for catalyticallyhydrodesulfurizing gas oil is admitted in pipe 4 at a pressure of 750p.s.i.g. The quantity and composition of this gas are specified incolumn 2 of Table I. This make-up gas joins the recycle gas stream,which is described later, in conduit 6. The resulting mixture has atemperature of 125 F. and its composition and rate of flow aredesignated in column 3 of the table.

Half of the mixed gas stream in conduit 6 is taken off in the valvedline 8 for purposes that will be apaprent later. The other half of thegaseous material continues to travel along pipe 6 until it joins thepyrolysis liquid hydrocarbons in conduit 2, and this gas-liquid mixtureof the composition and flow rate given in column 4 of Table I passesthrough the heater 10 Where its temperature is adjusted to 115 (hereindesignated as the feed temperature) by heating, if necessary, on its wayto reactor 12. This charge temperature produces good results with thecatalyst described hereinafter which has been partially deactivated inservice.

Column 4 of Table I sets forth the total charge to the first or initialreactor 12 which contains a fixed or stationary catalytic bed 13 ofchromia-promoted palladium on a gamma alumina support in the form of 7diameter cylinders long. Based on the total weight, there is a surfacedeposit on the alumina of 0.50 percent of palladium metal and also 0.51percent of chromium in the form of oxides,

The reaction conditions in the first reactor 12 are:

Inlet temperature F 115 Outlet temperature F 170 Pressure p.s.i.g 730Hydrogen partial pressure (inlet) p.s.i 450 H :liquid feed chargingratio s.c.f./B 1900 Liquid space velocity v./hr./v 2.2 Catalyst ActivityIndexes Hydrogenation 100 Polymerization 22 Desulfurization In the firststage reaction the primary reaction is one of the nondestructivehydrogenation of diolefins, especially conjugated diolefins, accompaniedby considerably less saturation of the less reactive mono-olefins. Thetemperatures are below the level required for desulfurization and nosignificant hydrogenation of aromatics or polymerization takes placethere.

Any trace of gum formed in the catalyst bed dissolves in the descendingliquid and the reaction efiluent is drawn off at the bottom of thereactor via conduit 14 in which it is transported to heater 15. A minorportion of the liquid feed stock or reaction products thereof vaporizesin reactor 12 as a result of the heat evolved in the exothermichydrogenation reaction.

A circulating flux liquid at 350 is injected from the conduit 16 intothe products in pipe 14 partly to increase the temperature of theinitial reactor effluent about 45 thus promoting its vaporization butchiefly to reduce any tendency toward the deposition of any gummy solidsin the transfer line 14. This flux liquid is drawn off near the bottomof the separator 18 in pipe 16 and recirculated by pump 22 at the rateof 9,220 lbs/hr. or 720 b./d.

This liquid is composed of the higher boiling hydrocarbons of theinitial reactor effiuent which are retained in the liquid phase and asmall quantity of dissolved polymeric material. The latter is aby-product of the present process and is readily soluble in the benzeneand other aromatic hydrocarbons constituting most of the liquid flux.

Two other modes of heating the first reaction effluent are also employedduring its passage to the vaporizer pot 18. Saturated steam at 220p.s.i.g. is admitted to the heater 15 under a control techniquedescribed hereinafter to indirectly heat the first reaction products toa temperature of 337. In addition, a heated hydrogen-rich gaseousmixture is injected into those initial products in conduit 20 upstreambut close to the chamber 18. This hydrogenrich stream is part of thatdrawn off in line 8 from the total circulating gas (recycle and make-upgases) in conduit 6. The gas in pipe 8 flows through the heat exchanger24 where its temperature is raised to 380 and finally into gas-firedheater 26. Firing of this heater is controlled in a unique manner whichis described later; and it provides an efiluent leaving in conduit 28 ata temperature of 645, which is divided by means of the three-way valve30 with 20% of the total circulating gas being introduced into pipe 32and the remaining 30% passing through conduit 34 to join the firstreaction efiluent in line 20. This further heating of the product streamin line 20 of course results in more vaporization and vaporization iscompleted to the desired extent in the flash chamber and separator 18.The latter is a vessel of enlarged cross section with an internaldiameter of 4.5 feet and a height of 12.5 feet which provides favorableconditions for the substantially complete separation of the gaseousphase from the liquid phase in a mixture thereof at a temperature of 360and pressure of 695 p.s.i.g.

Based on the rate of feeding pyrolysis gasoline, 4% of said liquid feedand reaction products thereof vaporizes in reactor 12, about 75% more isevaporated during passage through line 14 and heater 15, furthervaporization is produced by the hot gas injected from pipe 34 and only8.5% is collected in the liquid phase in the vaporizer pot 18 inaddition to the circulating flux.

The gaseous phase going overhead passes through the demister blanket orpad 36 of coarse steel wool designed to catch any entrained droplets ofliquid. No substantial deposition of polymers or gums occurs in thelines 14 and 20 or heater 15, but the liquid in the bottom of pot 18contains an amount of dissolved polymer (ASTM gum content=67 mg./ ml.)which is small but sufficient to foul and thereby deactivate a contactcatalyst within a fairly short time, particularly at desulfurizingtemperatures, of 450 and higher. A portion of the flux is continuallybeing removed at a constant rate of 200 b./d. as spent flux through thebottom line 38 under the regulation of the How controller 40 operatingthe automatic valve 42. The rate of withdrawing spent flux from thesystem is manually reset on that controller from time to time to theminimum rate that will hold the gum content thereof below about 100milligrams per 100 mls. The spent liquid flux is transferred to a reruntower (not shown).

While an extraneous flux may be alternatively supplied to the system ata constant rate through the line 50 connected to separator 18, asuitable flux is obtained from the efiiuent of the first reactor bytemporarily operating heater 15 in the manner described hereinbefore toaccumulate sufficient liquid in pot 18 for recycling as a flux; andthereafter normal operating conditions are employed in the vaporizingsystem. The overhead or vapor phase passes through heat exchanger 52 onits way from separator 18 via conduit 54 to join the hydrogen-rich gasfrom pipe 32 in line 56 as the charge for the second reactor 58. In thispassage, the heat exchanger 52 raises the temperature of the overheadeffiuent to 485 and adraises the temperature of the total charge to 515at the reactor inlet.

As indicated previously, two temperature control techsize prepared byhydrogen sulfide treatment in the manner described hereinbefore Withasulfur content of 4.6% at operating equilibrium and a weight ratio of AlO zMozCo of 84.7:7.9:2.7 respectively.

niques are employed for heating and thereby vaporizing The lessreactivediolefins remaining in the initial reac liquid efliuent from the firstreactor to prepare a vapor tion effiuent are saturated in the secondreactor along with phase charge for the second reactor. First, the rateof all of the mono olefins that remain'in a nondestructive flow ofheating steam through conduit 59 to heater 15 is manner and with no.substantial saturation ofaromatic controlled by automatic valve 60 inresponse to an eX- compounds. Most of this.addition of hydrogen tounsatu ternal liquid level controller 62 which is connected in ratedcompounds occurs inathe intermediate hydrogenaconventional manner tosense the liquid level in sepation (upper) zone. The reaction conditionsin the second. rator pot 18. Since the'rates of circulation of fluxliquid reactor are as follows: and remqvalof the spent u a ecustomarilyheld Inlet temperature degrees 7 515 stant, a rise in thelevel of liquid in pot 18 indicates that Average reaction temperature.the liquid feed stock and its liquid products are being Upper zone 71 do535 vaporized at a lower rate. This is corrected automatically Lowerzone .72 do 555. by the level controller 62 generating a function orsignal Outlet temperature, 555 in response to which valve 60automatically opens to adp Total pressure p.s.i.g 685 mit more steaminto heater 15 and thus vaporize more H partial pressure (inlet) p.s.i345 of the first reactor effiuent passing through the heater 15. 2O ITotal H charged s.c.f./b 3950- Conversely, a fall in liquid level in thevaporizing cham- V Space velocity-LHSV. ber indicates that a greaterproportion is being vaporized,

Upper zone 71 17 and this is corrected by a signal from the levelcontroller Lower zone 72 1.9 62 to the automatic valve 60 which reducesthe steam m-,

. Upper zone catalyst activity indexes. put to heater 15, and thereforeresults in a lower rate of I I Hydrogenation 100 vaporization in theliquid passing therethrough. v

Polymerization 22 The firing of the furnace 26 for heating hydrogen-richD If 0 circulating gas is controlled by the automaticvalve 64= esu unzaIon "."7"?

Lower zone catalyst activity indexes: operating in the fuel gas supplyline 66 in response to two D 100 esulfurization 98- temperaturecontrollers. Temperature controller 68 senses Pol merization 43 thetemperature in the outlet line 28 from the heater and maintains atemperature 645 at this point, but this de- From the inlet and outlettemperatures given, t is:apvice is reset to other temperatures as may berequired in Parent that Significant hydrogenation Ieacflons Wlth811bresponse to the temperature controller 70 which'is constflntifll'BXOthefmS taking Place in both reactQfsnected to conduit 56 andmaintains a temperature of'515 This is borne out by a comparison of theunsaturation in the charge entering the second reactor. indexes ofcolumn 4 with column 5 inT able I hereinafter Reactor 58 contains twobeds of contact catalysts. and also of column-6 With 7. The latter twoindicate Upper bed 71 made up of 1% platinum on chialumina that a minorhydrogenationof diolefins is completed in of 7 particle size occupies10% of the catalyst space the final reactor along with the principalhydrogen in the second reactor and the other 90% is filled with thetreatment that saturates substantially all of the remaining lower bed 72of desulfurization catalyst. A horizontal half of the mon o-olefins anda substantially complete perforate plate or screen (not shown) withopenings of hydrodesulfurization of; organic sulfur-compounds. /smaximum dimension is located at the boundary be- Again there is. noapprec able POlYIIlCI'lZdUOIl or depos tween the upper intermediatehydrogenation zoneand the tion of coke and no noticeable conversion ofaromatic lower desulfurization zone. This plate supports the upperhydrocarbons to naphthenes occurs.

Table l Stream Gasol. Fresh Total First First Final 1 Final H.P. Sep.Stabilizer Stabilized Feed H -Rich Hz-Rich Reactor Reactor ReactorReactor Ofi Gas Ofi Gas Liquid Gas Gas Charge Effluent Charge EfliluentI Flow, lbs./hou.r 28,220 2, 520 14, 020 35, 230 35,230 39,680 39,6801,860 920 25, 400 Gravity, API 36.8 37. 5 Ave. Moi. Wt 31. i 31. 2 80.1Bromine N0 24. 12 N l Diene Nor- 13. 5 1. 5 Nil Percent Vap d Nil 4Organic S, p.p.m. 450 450 10 ii fii 195. 0 961. 0 480. 5 79. 5 501. 2250. 6 20.1 34.2

3.5 31.9 30.3 277.2 3.6 15.3 Total, Mols/Hr 353.7 300.0 1, 560.9 1,134.2

1 Based on weight of original liquid feed.

bed and prevents any appreciable intermingling of the two catalysts. I

The lower bed is a composite of cobalt and molybdenum sulfides on agamma alumina of A inch particle 76, on its way to the second separator78 .where the vapor.

The gaseous product stream leaves the bottom of reactor 58 via conduit74' and is cooled by passing through heat exchangers 52 .and 24respectively, as Well as the cooler phase is separated from the newlycondensed liquid at a temperature of 100 and pressure of 640 p.s.i.g.From this vessel the gaseous phase is taken overhead in lines 80 and 82.About 15% of this gas is bled oif to the refinery fuel system throughpipe 84 and the pressure regulator 86 which maintains the desiredpressure on the hydrogenation system. The rate of removal of thisseparator gas from the instant system is tabulated in column 8 of TableI. Most of the gaseous material, however, enters the line 90 wherein itmeets with any make-up gas that also may require scrubbing to removeexcessive hydrogen sulfide which make-up gas is drawn from supplyconduit 4 via valved line 92. These gases are introduced into the lowerhalf of the combination washer 94 which is equipped with a lower causticscrubber section 96 beneath a water washing section 98.

Fresh aqueous sodium hydroxide solution is admitted in conduit 100 andjoins recirculating caustic soda solution in the line 102 on its way tothe perforated scrubber trays over which it cascades downwardly againstthe rising gases. This alkaline liquid is drawn off through the conduit106 at the bottom and divided between an exit line 108 for spentsolution and conduit 110 leading to the recirculation pump 112.

Cold water is admitted to the section 98 of the tower from supply line114 and is drawn oif through the valved conduit 116. It will be notedthat substantially all of the water is collected in the trough 118 andis not allowed 'to descend therebelow and dilute the caustic scrubbingsolution. The gases rising countercurrently through the tower 94 at 640p.s.i.g. lose most of their hydrogen sulfide content in being scrubbedfirst by intimate contact with curtains of caustic soda solution, nextthey pass through the demisting pad 120 into the washing section wherethey are washed with curtains of falling water to remove the last tracesof H 8 as well as any entrained particles of the caustic soda solutionand then through the demisting pad 122.

The scrubbed and washed gases exit through the conduit 124 whichconnects with the valved by-pass line 126, that may be used to divertsome or all of the separator off-gas around tower 94. The by-passconduit is useful when the hydrogen sulfide content of the separator gasis low enough for a recycle gas. These two pipes feed into the line 128which leads to the knockout pot 130 in which any entrained liquid isseparated. From here the hydrogen-rich gas passes through conduit 132 tocompressor 134 where its pressure is boosted sufficiently to circulateit through the recycle gas line 6 and associated conduits in the mannerdescribed earlier.

Returning now to the scrubber 94, it is apparent that an extremelyflexible arrangement is shown for controlling the hydrogen sulfidecontent of the circulating gases passed into the two reactors with thefeed. For example, the operator can divide the gaseous product fromseparator 78 between inlet line 90 of the caustic scrubber and theby-pass conduit 126 in any desired proportions. Similarly, the make-upgas entering in conduit 4 can be introduced directly into circulatinggas line 6 or part or all of it can be taken off via conduit 92 fortreatment in the caustic scrubber. Also, either or both of the rates ofrecirculation or caustic soda solutions in scrubber section 96 and theintroduction of fresh caustic soda thereto can be controlled to set therate of reaction and removal of hydrogen sulfide from the gas streampassing through the tower.

The liquid phase withdrawn from the bottom of high pressure separator 78is treated in the stabilizing tower 136 at 180 p.s.i.g. after beingcarried in the conduit 138 through the pressure reducing valve 140 andheat exchanger 142 in which the temperature of the stream is raised to240 F. Attached to the 30-tray stabilizer are valved inlet lines 144 to146 to introduce the charge selectively and in any proportions onto the18th and 12th trays respectively counting from the bottom of the column.A reboiler 148 is provided to maintain the bottoms at a temperature ofabout 385 F. and a stable, substantially saturated liquid product iswithdrawn as the product of the process via pipe 150 into heat exchanger142 at the rate given in column 10 of the table. This liquid, rich inaromatic hydrocarbons, is suitable for extraction processes, such asextraction with diethylene glycol, for the concentration of aromatics byreason of its negligible content of diolefins, olefins and sulfur. It isessentially a mixture of parafiinic and aromatic hydrocarbons, and asharp separation can readily be obtained between these constituents.

An overhead fraction is conveyed via the conduit 152 and cooler 154 inwhich cold water reduces its temperature from 295 to 125 in transit tothe reflux accumulator 156. Liquid reflux is returned from the bottom ofthis accumulator to the stabilizer 136 through line 158 and pump 160 ata rate of 9840 lbs. per hour and a gaseous by-product of the process iswithdrawn through the valved conduit 162 at the rate set forth in column9 of Table I for use as fuel or other suitable purposes.

Starting up the process described herein in a commercial plant isrelatively free of difliculties. Make-up gas obtained from a catalyticreformer is charged at ambient temperature and the usual operatingpartial pressure of hydrogen into the initial reactor 12 and alsothrough the furnace 26 into the final reactor 58. This is continueduntil the heat carried by the gas from the furnace brings the secondreactor close to its normal operating temperatures. Meanwhile, recyclegas is substituted for most of the fresh supply of hydrogen rich gasafter a thorough purging of the system. Next a typical reformate derivedfrom naphtha and relatively free from unsaturated aliphatic compounds isintroduced as a temporary feed along with the circulating gas. Anunusually high proportion of liquid accumulates in the separatingchamber 18 from the time the reformate is first charged until thatchamber reaches normal operating temperature and none is withdrawnthrough the spent flux line at first. After the circulating flux systemis allowed to fill up with the liquid phase collecting in the vaporizerchamber, liquid is drained off in the spent flux line at an abnormallyhigh rate until normal feed is being processed at normal vaporizingtemperatures. The final step is to gradually blend 'the regularpyrolysis liquid feed stock into the reformate in gradually increasingproportions with a corresponding reduction in the supply of the reformedproduct until the latter component is shut ofi completely.

EXAMPLE 2 Several considerably broader cuts of pyrolysis gasoline aresubjected to a multistage hydrogen treatment with a differentcombination of catalysts. A platinum catalyst is charged into theinitial reactor of a pilot plant. The upfiow second reactor contains abottom bed (9% of the total catalyst volume of the second reactor) ofthe same platinum catalyst used in the first reactor, and the upper bedconsists of 91 volume percent of 15.3% by weight of unsulfided cobaltmolybdate on gamma alumina. The gas charged is hydrogen of commercialpurity in place of the usual refinery and recycle gases containingsubstantial contents of lower hydrocarbons such as methane. In thesesmall scale operations, it is not feasible to recirculate the flux oreven remove the spent flux continuously; but highly colored spent fluxwith a significant gum content is withdrawn from the bottom of aseparator located between the two reactors at intervals of 8 hours inquantities equal to 0.5% by weight of the pyrolysis liquid charged. Theliquid phase of the initial reactor efiiuent 1s vaporized by injectinghot hydrogen into the aforesaid separator and also by controlledelectrical heating elements wrapped around the separator.

Other reaction conditions and the results obtainable during a lengthyoperation are set forth in Table II. No fouling from excessive formationof coke or polymers the examples, without departing from the invention.For

instance, standby units arranged in parallel with alternate piping maybe provided for all equipment that, requires periodic regeneration orlcle'aning. Accordingly, the pres? Table 11 1st Reactor 2nd ReactorReaction Conditions: 9 vol. percent (1% Pt/alumina) yst 9 1% Pt on F-chi alimina 91 vol. percent (Co-Mo/alumina) Total Pressure, p s r g. 450450 450 450 450 450 450 450 450 450 450 l 450 i Inlet H2, p.s.l 400 400400 400 400 400 330 330 330 330 330 330 H2 Charge, 5.0 f lb 1,500 1, 5001, 500 1, 500 1, 500 1, 500 2,000 addtional Space Velocity, LHS 2. 2. 2.2. 0 2. 2.0 20 in Pt Cat and 2.0 in Co-Mo Cat. Av. Reactor Temp, 225 225225 225 225 259 500 500 500 500 1 (500) Av. Temp, 2nd Pt Cat 520 516 528540 515 559 Cat. Age, Days 5. 3 11.1 18. 0 26; 1 32. 7 107' 5. 0 10. 818.0 26. 7 33.7 100 Charge Stock:

Bromine Number 39. 9 39. 9 43. 7 43. 7 25. 4 25. 8 Diene Number. 29 29'29 16 18 Sulfur, p.p.m 170 170 190 190 185 200 Initial B.P., I* 120 120142 161 End Point, F- 314 314 280 266 Product:

Bromine Number..- 17. 2 18. 3 25. 3 27. 4 16.0 22. 7 0.1 0.3 0. 3 0. 90.2 0. 3 Diene Number Nil 1 4 6 3 Nil Sulfur, p.p.m 160 160 9 5 9Catalyst Performance:

Percent Reduction, Br. No 57 54 42 37 37 12 Percent Reduction, DieneNo..." 100 97 87 r 79 81 28 1 Average Rea ction Temperature of only thecobalt molybdate catalyst bed.-

EXAMPLE 3 A different combination of catalysts is employed in theequipment used in Example 2 in hydrogenating another batch of pyrolysisgasoline to further illustrate the present invention. The intermediatehydrogenation zone: in the upper of the catalyst space is filled with a.

catalyst comprising 1% palladium on gamma alumina in the form ofextruded pellets. The balance of the catalyst volume contains gammaalumina bearing 3% cobalt oxide and 12% molybdenum oxide.

Significant reaction conditions and results are tabulated in Table III.

1 55% Ni on kieselguhr. 2 2(1)] vp}. percent Pd/alurnina, 80 vol.percent Co-Mo/alurnma. 3 0-1 0.

The selectively of the nickel catalyst for hydrogenating dienes inpreference to mono-olefins is apparent upon' comparing the Diene No. of1 and Bromine No. of 46 for the first stage product with the values of 8and 40 respectively that are obtained With the same feed and reac-- tionconditions except for substituting the 1% platinum catalyst of ExampleII.

The detailed examples given hereinbefore are intended only to illustratethe invention; It will be apparentto those skilled in the art that manyother modifications and variations may be made in the embodiments setforth in ent invention is not-to be considered as limited .in anyrespect other than the recitals of the appended claims.

Certain aspects of thereactions and/or the vaporization between reactionstages which are disclosed hereinbefore are described also and claimedinapplication Ser; No.

' 238,693 filed November 19, '1962 of Raymond R'.'Halik et al. entitled.Sele'ctive Hydrogenation of Hydrocarbons. and applicationSer.'No.-23,8,690* filedaNovember 19., 1962 of Richard GJGraven et a1.entitled ,Sele'ctive Conversion of Unstable? Liquids]? What is claimedis:

1. A process for the selective nondestructive hydro;

genation of a liquid hydrocarbon feed. boiling below about 500 F.' andcontaining ;aromatic hydrocarbons,

olefins, diolefins and sulfur compoundsjwhich.comprises passing saidfeed in the=liquid phase and hydrogen through an initial hydrogenationzone in contact with a porous solid hydrogenation catalyst having a highhydrogenation activity and a low polymerization activity whilecontrollinghydrogenating conditions in said. zone to provide a 1hydrogenation efiluentfrom said zone in: whichia substantial amountofthe diolefins have been atleast partially saturated and in which asubstantial part of said liquid feed and products thereof are in theliquid'phase, vaporizing. liquid in. said 1initial hydrogenationeflluent, passing the resulting vapors together with hydrogen through anintermediate hydrogenation Z0116? in contactwith a porous solidhydrogenation catalyst having a high hydrogenation activity and alowpolymerization activity" at a temperature substantially higher-thanthe average temperature in said initial zone under conditions controlledto. further hydrogenate said vapors; passing the efiluent from saidintermediate zone through a subsequentconversion zone at a suitable'desulfurization temperature .incontact with a porous solidsulfur-resistant conversion cata'lyst' having at least moderatehydrogenation activity and a high desulfurization activity, andregulating conditions in said conversion zone to ,producea substantiallydesulfurized conversion efiiuent with a normally liquid fraction havinga substantial lower Bromine Number-than said" liquid feed.

2. A .process for the selective nondestructive hydrogenation of a liquidhydrocarbon feed boiling below 25 about 500 F. and containing aromatichydrocarbons, olefins, diolefins and sulfur compounds which comprisespassing said feed in the liquid phase and hydrogen through an initialhydrogenation zone in contact with a porous solid hydrogenation catalysthaving a high hydrogenation activity and a low polymerization activitywhile controlling hydrogenating conditions in said zone to provide ahydrogenation effluent from said zone in which at least about 35% of thediolefins have been at least partially saturated and in which asubstantial part of said liquid feed and products thereof are in theliquid phase, vaporizing liquid in said initial hydrogenation effiuent,passing the resulting vapors together with hydrogen through anintermediate hydrogenation zone in contact with a porous solidhydrogenation catalyst having a high hydrogenation activity and a lowpolymerization activity at a temperature high enough for olefinsaturation and substantially higher than the average temperature in saidinitial zone under conditions controlled to further hydrogenate saidvapors whereby the diolefin content of the normally liquid fractionthereof is less than about 50% of that of said liquid feed, withdrawingthe efiluent from said intermediate zone at a temperature suitable fordesulfurization, passing said intermediate efiiuent through a subsequentconversion zone in contact with a porous solid sulfur-resistantconversion catalyst having at least moderate hydrogenation activity anda high desulfurization activity, and controlling conversion conditionsin said conversion zone to produce an effluent with a normally liquidfraction having a Bromine Number less than about 4 and an organic sulfurcontent below about 20 p.p.m.

3. A process according to claim 2 in which the temperature in saidintermediate zone is also high enough for desulfurization.

4. A process according to claim 2 in which at least about 50% of themore reactive diolefins are at least partially saturated in said initialzone and the diolefin content of the normally liquid fraction of saidintermediate efiiuent is less than about 40% of that of said liquidfeed.

5. A process according to claim 2 in which the charge to saidintermediate zone contains a sulfur compound in the gaseousphase havinga hydrogenation-inhibiting effect equivalent to 50 p.p.m. of thiophenesulfur and the partial pressure of hydrogen sulfide in said zone ismaintained below about 0.05 p.s.i.a., whereby conversion of aromatichydrocarbons into naphthenes is substantially prevented withoutsubstantially deactivating said intermediate hydrogenation catalyst.

6. A method according to claim 2 in which said initial and intermediatecatalysts have hydrogenation activity indexes of at least about 40,sulfur-free benzene conversion indexes above about 50 and polymerizationactivity indexes less than about 35; and said conversion catalyst has asulfur-free benzene conversion index below about 25, a freshdesulfurization activity index of at least about 80 and a polymerizationactivity index above about 25.

7. A method according to claim 2 in which said initial and intermediatecatalysts each contain a metal of Group VIII of the Periodic Table ofElements having an atomic number of at least 27, and said conversioncatalyst contains a metal of the iron group and a metal in Group VI B ofsaid Periodic Table.

8. A method according to claim 2 in which said initial and intermediatecatalysts comprise platinum supported on the surface of particle formalumina and said conversion catalyst comprises sulfides of cobalt andmolybdenum supported on the surface of particle form alumina.

9. A method according to claim 2 in which said initial catalystcomprises palladium supported on the surface of particle form alumina,said intermediate catalyst comprises platinum supported on the surfaceof particle form alumina said said conversion catalyst comprisessulfides of cobalt and molybdenum supported on the surface of particleform alumina.

10. A method according to claim 2 in which said initial catalystcomprises between about 0.05 and 10.0% palladium supported on thesurface of particle form alumina, said intermediate catalyst comprisesbetween about 0.05 and 2.0% platinum supported on the surface ofparticle form alumina and said conversion catalyst comprises compoundsof cobalt and molybdenum supported on the surface of particle formalumina.

11. A method according to claim 2 in which the temperature of saidintermediate efliuent is not substantially changed between saidintermediate zone and said conversion zone.

7 12. A method according to claim 2 in which said intermediate zone andsaid conversion zone are located in a single closed reaction vessel.

13. A method according to claim 2 in which said intermediate catalystand said conversion catalyst are employed in volumetric ratios between1:40 and 1:2, respectively.

14. A method according to claim 2 in which the conditions controlledwithin said initial zone include maintaining a hydrogen partial pressurewithin the range of about 200-800 p.s.i., an hourly space velocitywithin the range of about 0.2-15.0 based on the volume of said liquidfeed, a hydrogen charge within the range of about 500- 5000 s.c.f./b. ofsaid liquid feed and a feed temperature within the broad range of about75300 F., the conditions controlled in said intermediate zone includemaintaining a hydrogen partial pressure within the range of about200-800 p.s.i., an hourly space velocity within the range of about 2-60based on the volume of said liquid feed, a total hydrogen charge withinthe range of about 50010,000 s.c.f./ b. of said liquid feed and an inlettemperature within the wide range of about 350-700 F.; and theconditions controlled in said conversion zone include maintaining anhourly space velocity between about 0.2 and 8 based on the volume ofsaid liquid feed and an average reaction temperature not substantiallybelow said inlet temperature of the intermediate zone.

15. A method according to claim 14 in which said feed temperature ismaintained within the narrow range of 75190 F. while said catalyst isfresh and said temperatuer is increased within the limits of said broadrange to maintain said diolefin saturation as the hydrogenation activityof said initial catalyst decreases with continued use, and said inlettemperature to the intermediate zone is maintained within a narrow rangeof about 400550 F. while said intermediate and conversion catalysts arefresh and said inlet temperature is increased within said wide range asthe activity of said catalysts decreases with continued use in order tomaintain said organic sulfur content and Bromine Number in saidconversion efiluent fraction.

16. A method according to claim 2 in which the conditions controlledwithin said initial zone include maintainmg a hydrogen partial pressurewithin the range of about 300600 p.s.i., an hourly space velocity withinthe range of about 0.5-8.0 based on the volume of said liquid feed, ahydrogen charge within the range of about 12003000 s.c.f/ b. of saidliquid feed and a feed temperature within the range of about 75250 F. toprovide a hydrogenation efliuent from said zone in which the BromineNumber of the normally liquid fraction thereof is at least 25% belowthat of the liquid feed and at least about 50% of the diolefins havebeen at least partially saturated and in which an amount equal to atleast about 60% of the liquid feed is in the liquid phase; theconditions controlled in said intermediate zone include maintaining ahydrogen patrial pressure within the range of about 300-600 p.s.i., anhourly space velocity within the range of about 5-40 based on the volumeof said liquid feed, a total hydrogen charge within the range of about20005000 s.c.f./b. of said liquid feed and an inlet temperature withinthe range of about 400650 F. to provide an intermediate zone efiluentwith a normally liquid fraction in which the References. Cited by: theExaminerdiolefin content is less than about 40% of that of said I liquidfeed; and the conditions controlled in said con- UNITED STATES PATENTSversion zone include maintaining an hourly space velocity 2,901,4178/1959 Cook 'etzal; 208-210 between about 0.5-5 based on the volume ofsaid liquid 5 3,025,230 3/ 1962 l MaCLareIl t 208'210 feed and anaverage reaction temperature not substantial- 3,077,448v 1 Kal'dash 31;7 '1y below said inlet temperature of the intermediate zone 3,119,7651/4964 COmeil 208-;210

to produce an efiluent with a normally liquid fraction having a BromineNumber less than about 2.0 and an or- DELBERT GANTZ P'lmary Examinerganic sulfur content below about 15 p.p.m. 10 ALPHONSO D. SULLIVAN,Examiner.

1. A PROCESS FOR THE SELECTIVE NONDESTRUCTIVE HYDROGENATION OF A LIQUIDHYDROCARBON FEED BOILING BELOW ABOUT 500*F. AND CONTAINING AROMATICHYDROCARBONS, OLEFINS, DIOLEFINS AND SULFUR COMPOUNDS WHICH COMPRISESPASSING SAID FEED IN THE LIQUID PHASE AND HYDROGEN THROUGH AN INITIALHYDROGENATION ZONE IN CONTACT WITH A POROUS SOLID HYDROENATION CATALYSTHAVING A HIGH HYDROGENATION ACTIVITY AND A LOW POLYMERIZATION ACTIVITYWHILE CONTROLLING HYDROGENATING CONDITIONS IN SAID ZONE TO PROVIDE AHYDROGENATION EFFLUENT FROMSAID ZONE IN WHICH A SUBSTANTIAL AMOUNT OFTHE DIOLEFINS HAVE BEEN AT LEAST PARTIALLY SATURATED AND IN WHICH ASUBSTANTIAL PART OF SAID LIQUID FEED AND PRODUCTS THEREOF ARE IN THELIQUID PHASE, VAPORIZING LIQUID IN SAID INITIAL HYDROGENATION EFFLUENT,PASSING THE RESULTING VAPORS TOGETHER WITH HYDROGEN THROUGH ANINTERMEDIATE HYDROGENATION ZONE IN CONTACT WITH A POROUS SOLIDHYDROGENATION CATALYST HAVING A HIGH HYDROGENATION ACTIVITY AND A LOWPOLYMERIZATION ACTIVITY AT A TEMPERATURE SUBSTANTIALLY HIGHER THAN THEAVERAGE TEMPERATURE IN SAID INITIAL ZONE UNDER CONDITIONS CONTROLLED TOFURTHER HYDROGENATE SAID VAPORS, PASSING THE EFFLUENT FROM SAIDINTERMEDIATE ZONE THROUGH A SUBSEQUENT CONVERSION ZONE AT A SUITABLEDESULFURIZATION TEMPERATURE IN CONTACT WITH A POROUS SOLIDSULFUR-RESISTANT CONVERSION CATALYST HAVING AT LEAST MODERATEHYDROGENATION ACTIVITY AND A HIGH DESULFURIZATION ACTIVITY,ANDREGULATING CONDITIONS IN SAID CONVERSION ZONE TO PRODUCE A SUBSTANTIALLYDESULFURIZED CONVERSION EFFLUENT WITH A NORMALLY LIQUID FRACTION HAVINGA SUBSTANTIAL LOWER BROMINE NUMBER THAN SAID LIQUID FEED.